Process for converting oxygenates to alkylated liquid hydrocarbons

ABSTRACT

Alkylate is produced by catalytically converting oxygenate feedstock, such as methanol, to lower olefins comprising C 2  -C 4  olefins. Ethene is separated by interstage sorption of C 3  + components and an isoparaffin is alkylated with C 3  -C 4  olefins derived from sorbate. 
     The improved technique comprises fractionating an olefinic feedstream containing ethene and C 3  + olefinic components by contacting the olefinic feedstream in a sorption zone with a liquid hydrocarbon sorbent to selectively sorb C 3  + components; reacting C 3  + olefins with excess isoparaffin in a catalytic alkylation reactor to produce C 7  + alkylate hydrocarbons; fractionating the alkylation reactor effluent to provide a liquid hydrocarbon fraction rich in C 7  +, alkylate for recycle to the sorption zone as lean sorbent; and recovering C 7  + alkylate product and C 5  + gasoline from the process.

BACKGROUND OF THE INVENTION

This invention relates to an integrated system for convertingoxygenates, such as methanol or dimethyl ether (DME). to liquidhydrocarbons. In particular it provides a continuous process forproducing C₇ ³⁰ hydrocarbon products by converting the oxygenatefeedstock catalytically to an intermediate lower olefinic stream andalkylating isobutane or other isoparaffins with olefins to produce lightdistillate and/or gasoline products.

In order to provide an adequate supply of liquid hydrocarbons for use assynfuels or chemical feedstocks, various processes have been developedfor converting coal and natural gas to gasoline and distillate. Asubstantial body of technology has grown to provide oxygenatedintermediates, especially methanol. Large scale plants can convertmethanol or similar aliphatic oxygenates to liquid fuels, especiallygasoline. However, the demand for heavier hydrocarbons has led to thedevelopment of processes for making diesel fuel by a multi-stagetechnique.

Recent developments in zeolite catalysts and hydrocarbon conversionprocesses have created interest in utilizing olefinic feedstocks, forproducing C₅ ⁺ gasoline, diesel fuel, etc. In addition to the basic workderived from ZSM-5 type zeolite catalysts, a number of discoveries havecontributed to the development of new industrial processes.

The medium pore ZSM-5 type catalysts are useful for converting methanoland other lower aliphatic alcohols or corresponding ethers to olefins.Particular interest has been directed to a catalytic process forconverting low cost methanol to valuable hydrocarbons rich in ethene andC₃ ⁺ alkenes. Various processes are described in U.S. Pat. Nos.3,894,107 (Butter et al), 3,928,483 (Chang et al), 4,025,571 (Lago),4,423,274 (Daviduk et al), 4,433,189 (Young), and 4,543,435 (Gould andTabak), incorporated herein by reference. It is generally known that theMTO process can be optimized to produce a major fraction of C₂ -C₄olefins. Prior process proposals have included a separation section torecover ethene and other gases from byproduct water and C₅ ⁺ hydrocarbonliquids.

SUMMARY OF THE INVENTION

It has been discovered that methanol, DME or the like may be convertedto liquid fuels, particularly alkylate, in a multi-stage continuousprocess, with integration between the major process units providing anethene-rich recycle stream for further conversion and an alkylaterecycle stream for interstage sorption. The initial stage MTO typeprocess hydrocarbon effluent stream, after byproduct water separation,can be fed to an alkylation stage for conversion to heavierhydrocarbons. Ethene may be recovered by interstage separation andrecycled. Advantageously, the recycled ethene is found to be reactivewith methanol/DME or other oxygenates in the presence of ZSM-5 typecatalysts. In effect a novel MTO-Alkylation system is provided whereinthe ethene component may be recycled sustantially to extinction.

In a preferred embodiment, the invention provides processes andapparatus for an integrated continuous technique for convertingoxygenated organic feedstock to liquid alkylate hydrocarbons comprisingmethods and means for

(a) contacting feedstock with zeolite catalyst in a primary catalyststage at elevated temperature to convert at least a portion of thefeedstock oxygenate predominantly C₂ -C₄ olefins and a minor fractioncontaining C₅ ⁺ hydrocarbons;

(b) cooling and separating effluent from step (a) to provide an aqueousliquid byproduct stream, a heavy hydrocarbon liquid stream and a lighthydrocarbon vapor stream rich in C₂ -C₄ olefins;

(c) compressing at least a portion of the olefinic light hydrocarbonstream to condense a liquid olefinic hydrocarbon stream rich in C₃ -C₄olefins and recovering an ethene-rich gaseous stream;

(d) contacting the light hydrocarbon vapor stream from step (c) with aliquid hydrocarbon sorbent stream in a sorption tower under conditionsto selectively sorb the major amount of C₃ + hydrocarbon components fromsaid light vapor stream to provide a sorbate stream rich in C₃ -C₄olefins.

(e) further reacting the condensed liquid olefinic hydrocarbon streamfrom step (c) and the sorbate stream from step (d) with isoparaffin in asecondary alkylation stage with liquid phase acid catalyst to convert atleast a portion of C₃ -C₄ olefins to a heavier C₇ ⁺ liquid hydrocarbonproduct stream comprising alkylate gasoline;

(f) recycling ethene in a gaseous stream to the primary catalytic stage;and

(g) recycling a portion of alkylate gasoline to step (d) for use as leansorbent.

Advantageously, the primary stage catalyst comprises ZSM-5 type zeoliteand ethene is recycled to the primary stage at a rate of about 1 to 20parts ethene per 100 parts by weight of methanol equivalent in thefeedstock. By fractionating gaseous effluent separated from the primarystaged effluent, a recycle gas stream may be recovered containing atleast 90% of ethene from the primary catalytic stage. An olefinic streamrich in C₃ ⁺ olefins, especially propene and butylenes, is provided forreaction with various isoparaffins, such as isobutane.

Other objects and features of the invention will be seen in thefollowing description and drawings.

THE DRAWINGS

FIG. 1 is a process flow sheet showing the major unit operations andprocess streams;

FIG. 2 is a schematic representation of a preferred inter-stageseparation system for ethene recovery; and

FIG. 3 is an alternative process flow sheet depicting sorption,fractionation and alkylation units schematically.

DESCRIPTION OF PREFERRED EMBODIMENTS

Numerous oxygenated organic compounds may be contained in the feedstockmaterial to be converted in the primary stage. Since methanol or itsether derivative (DME) are industrial commodities available fromsynthesis gas or the like, these materials are utilized in thedescription herein as preferred starting materials. It is understood bythose skilled in the art that MTO-type processes can employ methanol,dimethylether and mixtures thereof, as well as other aliphatic alcohols,ethers, ketones and/or aldehydes. It is known in the art to partiallyconvert oxygenates by dehydration, as in the catalytic reaction ofmethanol over gamma-alumina to produce DME intermediate. Typically, anequilibrium mixture (CH₃ OH+CH₃ OCH₃ +H₂ O) is produced by partialdehydration. This reaction takes place in either conversion of methanolto lower olefins (MTO) or methanol to gasoline (MTG).

The zeolite catalysts preferred for use herein include the crystallinealuminosilicate zeolites having a silica to alumina ratio of at least12, a constraint index of about 1 to 12 and acid cracking activity ofabout 1-200. Representative of the ZSM-5 type zeolites are ZSM-5,ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-35 and ZSM-38. ZSM-5 is disclosedand claims in U.S. Pat. No. 3,702,886 and U.S. Pat. No. Re. 29,948;ZSM-11 is disclosed and claimed in U.S. Pat. No. 3,709,979. Also, seeU.S. Pat. No. 3,832,449 for ZSM-12; U.S. Pat. No. 4,076,979. Also, seeU.S. Pat. No. 3,832,449 for ZSM-12; U.S. Pat. No. 4,076,842 for ZSM-23;U.S. Pat. No. 4,016,245 for ZSM-35 and U.S. Pat. No. 4,046,839 forZSM-38. The disclosures of these patents are incorporated herein byreference. A suitable catalyst for oxygenate conversion is HZSM-5zeolite with alumina binder. These medium pore shape selective catalystsare sometimes known as porotectosilicates or "pentasil" catalysts.

Other catalysts and processes suitable for converting methanol/DME tolower olefins are disclosed in U.S. Pat. No. 4,393,265 (Bonifaz), U.S.Pat. No. 4,387,263 (Vogt et al.) and European Patent Application No.0081683 (Marosi et al.), and ZSM-45. In addition to the preferredaluminosilicates, the borosilicate, ferrosilicate and "silicalite"materials may be employed. ZSM-5 type catalysts are particularlyadvantageous because the same material may be employed for dehydrationof methanol to DME, conversion to lower olefins and ethylene conversion.

In this description, metric units and parts by weight are employedunless otherwise stated. Various reactor configurations may be used,including fluidized bed catalytic reactors, moving bed and fixed bedreactors.

Referring to FIG. 1, the process feedstock (methanol or DME, forinstance) is fed to the primary MTO stage (I) where it is converted tolower olefins and C₅ + gasoline hydrocarbon plus water by dehydration ofthe oxygenated feedstock. Byproduct water is recovered by simple phaseseparation from the cooled effluent. Liquid hydrocarbons consistingessentially of C₅ ⁺ gasoline range materials may be recovered or pumpedto the higher secondary stage pressure. Vapor phase effluent from theprimary stage may be compressed to alkylation reaction pressure.Propylene, butylenes and amylenes may be separated from the primarystage effluent by sorption fractionation to recover a recycle gas streamcontaining at least 90% of ethene from the primary stage and an olefinicsorbate stream rich in C₃ + olefins. A C₃ -C₄ rich olefinic stream maybe further prepared for reaction with isobutane or the like at highpressure and low temperature in contact with liquid phase acidicalkylation catalyst. Secondary stage (II) alkylation effluent is thenseparated into C₂ ⁻ light gases, C₃ -C₄ aliphatics and C₅ ⁺ gasolineand/or light distillate range hydrocarbons. Advantageously, isobutane isseparated from the second stage effluent for recycle to provide astoichiometric excess and added with fresh feed (i-C₄) to the unit. Aportion of the liquid alkylate-rich hydrocarbon is recycled to theinterstage sorption unit as lean sorbent. The preferred sorption unit isa countercurrent packed tower. Advantageously, the sorbent liquidcomprises at least 75 wt% of C₇ to C₉ isoparaffins which are thealkylate reaction product. This lean sorbent has excellent propertiesfor selective sorption of the propene, butylene and C₃ + paraffiniccomponents of the primary stage light hydrocarbons. Unfractionatedliquid alkylation effluent may be recycled in part, as depicted bydashed line.

The process may be optimized by employing fluid bed primary stageconditions in the temperature range of about 425° C. to 550° C., apressure range of about 100 to 800 kPa and weight hourly space velocityrange of about 0.5 to 3.0 based on ZSM-5 equivalent catalyst andmethanol equivalent in the primary stage feedstock. Suitable equipmentand operating conditions are described in U.S. patent application Ser.No. 687,045, filed Dec. 28, 1984, incorporated herein by reference.

In the embodiment of FIG. 2, the light hydrocarbon vapor streamseparated from the primary stage effluent is compressed in a pluralityof compression stages to condense liquid olefinic hydrocarbons. The fullreaction effluent of the primary stage MTO plant is passed via conduit201 and primary phase separator 202 to provide a first vapor stream richin C₄ -hydrocarbons, liquid hydrocarbons stream, and byproduct waterstream. The liquid (eq-C₅ ⁺) stream is combined with a correspondingliquid HC from succeeding separators and withdrawn. The primary vaporstream is adiabatically compressed by multi-stage compressor 205, cooledvia exchanger 206 and passed to a succeeding separator 204A, at whichpoint the preceeding phase separation technique is repeated. Likewiseother separators 204B and 204C operate to provide an ethene-rich recyclestream which is passed via line 214 to turbo-expander 209 and thus atMTO pressure back via line 211 to the olefins production in the primarystage. Advantageously, the MTO effluent is received at about atmosphericpressure (eg, 100∝150 kPa) and compressed in plural stages to a pressureof about 1100-3500 kPa (150-400 psig) and separated in the final vessel204C at about ambient temperature (20°-80° C.). Olefinic liquids rich inC₃ ⁺ aliphatic are recovered from the final compressor stage via pump208 which passes the liquid hydrocarbon stream to the followingsecondary stage alkylation unit.

Ethene-rich vapor withdrawn from the separator 204C via line 213 iscooled by heat exchange and further processed to increase ethene purityin sorption unit 216. A suitable selective sorption unit is disclosed inU.S. Pat. No. 4,497,968 (Hsia et al), incorporated herein by reference.Preferably, compressed light hydrocarbons are fractionated by sorptionto recover a recycle stream containing at least 90 mole percent ethene.This can be achieved by selectively absorbing C₃ ⁺ components in a C₅ ⁺liquid hydrocarbon sorbent stream, especially C₇ to C₉ alkylate.

Interstage fractionation can be modified within the inventive concept toprovide for recovery of purified ethylene (C₂ H₂) and heavierhydrocarbon liquids. In FIG. 3, a sorption fractionator unit 318,comprising a vertical countercurrent contact tower, is equipped with abottom reboiler means 319. Sorbent liquid, rich in C₇ to C₉ alkylate, isintroduced via conduit 320. Optionally, other C₅ + liquids may beemployed to supplement the alkylate stream. A light olefin feedstreamcontaining C₂ -C₄ alkenes is introduced via vapor conduit V and liquidconduit L. The sorber overhead may contain 40% or more ethene, which canbe further purified by cryogenic separation unit 330 to remove anyremaining C₃ + hydrocarbons and thus to recover pure ethene. The sorbatestream 325, rich in C₃ -C₄ alkenes, is fractionated in distillationtower 340 to recover gasoline range hydrocarbons, this tower may beoperated to obtain a heavy C₆ product, with C₅ components passing eitheroverhead or with the bottoms. Optionally, this overhead fraction isde-ethanized by tower 345, and the liquified sorbate fraction, rich inC₃ -C₄ components is combined with C₃ + components from the cryogenicseparation unit and fed at high pressure (e.g., up to 3000 kPa) toalkylation reactor unit 350, where it is contacted with an isoparaffin,such as isobutane (i-C₄) in the presence of an alkylation catalyst toproduce alkylate hydrocarbons. The alkylation fractionation system 360may be operated in a known manner to separate the reactor effluent intoseveral fractions, including i-C₄ recycle to provide excess isoparaffin,light gas (C₃), normal paraffin (n-C₄) and C₅ + liquid. At least aportion of the C₅ + components, especially the C₇ to C₉ alkylatereaction products, is passed via conduit 320 to the sorptionfractionator unit as lean sorbent. Recovered C₅ + liquids fromfractionation unit 360 may be further refined to recover aviationgasoline, light distillate, etc. If additional sorbent liquid isrequired, a bottom fraction from unit 340 may be utilized, as indicatedby dashed line 342. However, it is preferred to employ sorbentcomprising a major amount of C₇ -C₉ hydrocarbons, particularly at least75% paraffinic alkylate.

The data in Table I represents a material balance and absorber operatingconditions for a countercurrent contact tower unit design according toFIG. 3 and Runs 1-3. The vertical tower has 21 theoretical stages, withC₇ -C₉ alkylate (lean oil #1) from second stage product fractionatorbeing introduced at the top (stage 1), heavy liquid separated from theMTO effluent (lean oil #2) being introduced at stage 5, olefin vapor andliquid feed being fed at stages 12 and 13 respectively. Heat of sorptionis removed by pumparound cooling at stages 5 and 8. The three runscorrespond to different lean oil rates. In the overhead stream, molarunits are gm-moles per metric tonne (MT) of methanol (MeOH) charged tothe process. The operating conditions are chosen to provide a maximumethene content of 0.2 mol % in the sorbate.

                  TABLE 1                                                         ______________________________________                                        ABSORBER OPERATION*                                                           Run No.        1         2         3                                          ______________________________________                                         Material Balance                                                             wt % of MeOH charge                                                           (1) C.sub.6 + MTO Gasoline                                                                   50.0      22.0      17.6                                       (2) Alkylate   15.3      47.3      51.7                                       (3) n-butane   4.9       1.0       0.8                                        (4) propane    3.9       3.8       3.8                                        (5) offgas     1.5       1.5       1.6                                        (6) water      56.4      56.4      56.4                                       (7) i-C.sub.4 makeup                                                                         -32.0     -32.0     -31.9                                      (8) methanol   -100.0    -100.0    -100.0                                     Absorber*                                                                     Lean Oil #1 (stage 1)                                                                        4.44      1.19      0.69                                       g-moles/tonne MeOH                                                            Lean Oil #2 (stage 5)                                                                        1.01      1.01      1.01                                       g-moles/tonne MeOH                                                            TOTAL          5.45      2.20      1.70                                       overhead flow, SCM/ton                                                                       42.0      45.3      49.5                                       MeOH                                                                          propane, g-moles/tonne                                                                       0.0       6.5       18.5                                       MeOH                                                                          propylene      1.0       120.4     280.0                                      n-butane       3.9       18.3      18.1                                       isobutane      0.1       0.6       1.5                                        1-butylene     0.0       1.0       5.4                                        n-pentanes     0.7       0.6       0.6                                        iso-pentanes   40.2      30.0      28.4                                       pentanes       0.0       0.0       0.1                                        C.sub.6 +      16.5      17.4      17.9                                       Overhead pressure, kPa                                                                       2068      2068      2068                                       Cooling pump around,                                                                         -36.7     -35.4     -31.7                                      MJ/ tonne MeOH                                                                Reboiler Duty, MS/tonne                                                                      206.1     106.6     94.2                                       MeOH                                                                          Ethylene recovery in                                                                         938       938       938                                        overhead g-moles/                                                             tonne MeOH                                                                    Mole % purity  53        49        45                                         ______________________________________                                         *based on 0.2 mol % C.sub.2 in bottoms                                   

Runs 4 to 6 (Table II) are similar to Runs 1 to 3, except that theparaffinic lean oil is alkylation reactor effluent containing 78 mole %isobutane, 6% C₇ -C₉ alkylate and C₅ ⁻ hydrocarbons.

                  TABLE II                                                        ______________________________________                                        ABSORBER OPERATION                                                            Run No.        4         5         6                                          ______________________________________                                        Material Balance                                                              wt % of MeOH charge                                                           (1) C.sub.6 + MTO gasoline                                                                   15.0      12.3      11.9                                       (2) Alkylate   53.6      57.2      57.5                                       (3) n-butane   0.8       0.7       0.7                                        (4) propane    4.8       4.0       3.9                                        (5) offgas     1.5       1.5       1.6                                        (6) water      56.4      56.4      56.4                                       (7) iC.sub.4   -32.1     -32.1     -32.0                                      (8) methanol   -100.0    -100.0    -100.0                                     Absorber Efficiency*                                                          Lean Oil #1 (stage 5)                                                                        4.44      1.09      0.54                                       g-moles/tonne Meort                                                           Lean Oil #2 (stage 1)                                                                        1.01      1.01      1.01                                       g-moles/ton MeOH                                                              Total          5.45      2.10      1.55                                       Overhead flow, SCM/ton                                                                       45.8      46.7      50.3                                       MeOH                                                                          propene, g-moles/                                                                            47.7      37.2      36.4                                       tonne MeOH                                                                    propylene      36.1      97.2      252.1                                      n-butane       3.3       2.7       2.5                                        isobutane      62.7      47.7      39.7                                       i-butylene     34.9      35.7      36.5                                       n-pentanes     0.6       0.6       0.6                                        iso-pentanes   4.9       5.0       5.0                                        pentenes       13.0      13.3      13.5                                       C.sub.6 +      12.8      13.1      13.1                                       Overhead pressure, kPa                                                                       2068      2068      2068                                       Cooling Pump around,                                                                         -37.7     -27.9     -21.0                                      MJ/tonne MeOH                                                                 Reboiler Duty, MJ/                                                                           152.2     95.2      85.6                                       tonne MeOH                                                                    Ethylene recovery in                                                          overhead g-moles/tonne                                                        MeOH           939       939       939                                        Mole % purity  49        48        44                                         ______________________________________                                         *based on 0.2 mol % C.sub.2 in bottoms                                   

The alkylation process employed herein is a well known industrialtechnique for reacting alkenes with tertiary alkanes (isoparaffins),such as isobutane, isopentane, isohexane, etc. The resulting product isa C₇ ⁺ branched chain paraffinic material useful as aviation gasoline,jet fuel or the like. The alkylation of paraffins can be carried outeither thermally or catalytically; however, acid catalyst is preferred.Thermal or noncatalytic alkylation of a paraffin with an olefin iscarried out at high temperatures (about 500° C.) and pressures 21-41 MPa(3000-6000 psi). Under these conditions, both normal and isoparaffinscan be brought into reaction by a free-radical mechanism. Thermalalkylation is not known to be practiced commercially.

The catalytic alkylation of paraffins involves the addition of anisoparaffin containing a tertiary hydrogen to an olefin. The process isused in the petroleum industry to prepare highly branched paraffinsmainly in the C₇ to C₉ range, that are high-quality fuels. The overallprocess is complex, requiring control of operating conditions and ofcatalyst. The process conditions and the product composition depend onthe particular hydrocarbons involved.

The preferred processes are those brought about by the conventionalprotonic and Lewis catalysts. Propene can be brought into reaction withan isoparaffin in the presence of either concentrated sulfuric acid orhydrogen fluoride. The heptanes produced by alkylation of isobutane withpropene are mainly 2,3- and 2,4-dimethyloentane. Propene is alkylatedpreferrably as a component of a C₃ -C₄ fraction. HF catalysts foralkylation of isobutane with 1- and 2-butenes give both dimethylhexanesand trimethylpentanes. The product obtained from alkylation of isobutanewith isobutylene at low temperature (e.g., -25° C.) with hydrogenfluoride is 2,2,4-trimethylpentane.

During use the acid catalysts may become diluted with byproducthydrocarbons and as a result decrease in activity. Sulfuric acidconcentrations are maintained at about 90%. Hydrogen fluorideconcentrations of 80-90% are common, although the optimum concentrationdepends on the reaction temperature and reactor geometry. Operationbelow these acid concentrations generally causes incomplete conversionor polymerization. With sulfuric acid, the product quality is improvedwhen temperatures are reduced to the range of 0°-10° C. Coolingrequirements are obtained by low temperature flashing of unreactedisobutane. With hydrogen fluoride, the reaction process is lesssensitive to temperature, and temperatures of 0°-40° C. can be used.Some form of heat removal is essential because the heat of reaction isapproximately 14×10⁵ J/kg (600 Btu/lb) of butenes converted. Typicallythe elevated pressure for alkylation by these acid catalysts is about1500 to 3000 kPa (200-300 psig).

In order to prevent polymerization of the olefin as charged, an excessof isobutane is present in the reaction zone. Isobutane-to-olefin molarratios of 6:1 to 14:1 are common, more effective suppression of sidereactions being produced by the higher ratios.

The typical alkylation reaction employs a two-phase system with a lowsolubility of the isobutane in the catalyst phase. In order to ensureintimate contact of reactants and catalyst, efficient mixing isprovided. This is important with sulfuric acid because of the lowsolubility of isobutane in the catalyst phase. In addition, the higherviscosity of the sulfuric acid requires a greater mixing energy toassure good contact. The solubility of the hydrocarbon reactants in thecatalyst phase is increased by the presence of the unsaturated organicdiluent held by the acid catalyst. This organic diluent also has beenconsidered a source of carbonium ions that promote the alkylationreaction.

For the hydrofluoric acid system, reactive i-C₄ H₈ readily alkylates togive an excellent product. The alkylation of pure 1-C₄ H₈ by itselfproceeds with considerable isomerization of the 1-C₄ H₈ to 2-C₄ H₈followed by alkylation to give a highly branched product. The presenceof i-C₄ H₈ accelerates the alkylation reaction and allows less time forolefin isomerization. Consequently the reaction produces an alkylatewith a lowered antiknock value. For the sulfuric acid system, i-C₄ H₈tends to oligomerize and causes other side reaction products of inferiorquality; but the isomerization of 1-C₄ H₈ to 2-C₄ H₈ proceeds morecompletely, thereby favoring formation of superior products. Thus formixed olefin feeds such as described above, the two factors with bothcatalyst systems counteract each other to provide products of similarantiknock properties.

The olefin-producing MTO process may simultaneously generate isobutane,but the amount may be insufficient to alkylate the coproduced olefins. Asuitable outside source of isobutane is natural gas or a byproduct ofmethanol-to-gasoline (MTG) processes.

Suitable alkylation processes are described in U.S. Pat. Nos. 3,879,489(Yurchak et al), 4,115,471 (Kesler), 4,377,721 (Chester) and in theKirk-Othmer Encyclopedia of Chemical Technology, Vol. 2, pp. 50-58 (3rdEd., 1978) John Wiley & Sons, incorporated herein by reference.

The combined processes are an effective means for converting oxygenatedorganic compounds, such as methanol, DME, lower aliphatic ketones,aldehydes, esters, etc, to valuable hydrocarbon products. Thermalintegration is achieved by employing heat exchangers between variousprocess streams, towers, absorbers, etc.

Various modifications can be made to the system, especially in thechoice of equipment and non-critical processing steps. While theinvention has been described by specific examples, there is no intent tolimit the inventive concept as set forth in the following claims.

We claim:
 1. An integrated continuous process for converting oxygenatedorganic feedstock to liquid hydrocarbons comprising the steps of(a)contacting feedstock with zeolite catalyst in a primary catalyst stageat elevated temperature to convert at least a portion of the feedstockoxygenate predominantly to C₂ -C₄ olefins and a minor fractioncontaining C₅ ⁺ hydrocarbons; (b) cooling and separating effluent fromstep (a) to provide an aqueous liquid byproduct stream, a heavyhydrocarbon liquid stream and a light hydrocarbon vapor stream rich inC₂ -C₄ olefins; (c) compressing at least a portion of the olefinic lighthydrocarbon stream to condense a C₃ ⁺ liquid olefinic hydrocarbon streamrich in C₃ -C₄ olefins and recovering an ethene-rich gaseous stream; (d)contacting the light hydrocarbon vapor stream from step (c) with aliquid hydrocarbon sorbent stream in a sorption tower under conditionsto selectively sorb the major amount of C₃ + hydrocarbon components fromsaid light vapor stream to provide a C₃ ⁺ sorbate stream rich in C₃ -C₄olefins. (e) fractionating C₃ ⁺ sorbate to provide C₅ ⁺ gasoline and aC₃ -C₄ olefin stream; (f) further reacting the C₃ -C₄ olefin stream fromstep (e) with isoparaffin in a secondary alkylation stage with liquidphase acid catalyst to convert at least a portion of C₃ -C₄ olefins to aheavier C₇ ⁺ liquid hydrocarbon product stream comprising alkylategasoline; (g) recycling ethene for further catalytic conversion in theprimary catalytic stage; and (h) recycling a portion of alkylategasoline to step (d) for use as lean sorbent.
 2. The process of claim 1further comprising the step of fractionating gaseous effluent fromseparation step (b) to recover a recycle gas stream containing at least90% of ethene from the primary catalytic stage and an olefinic streamrich in C₃ ⁺ olefins.
 3. The process of claim 1 wherein the primarystage catalyst comprises ZSM-5 type zeolite and ethene is recycled tothe primary stage at a rate of about 1 to 20 parts ethene per 100 partsby weight of methanol equivalent in the feedstock.
 4. The process ofclaim 1 wherein primary stage feedstock comprising methanol and/ordimethyl ether and recycled ethene are converted over HZSM-5 catalyst toprovide a light olefinic hydrocarbon vapor stream comprising a majoramount of C₃ -C₄ olefins and a minor amount of ethene.
 5. The process ofclaim 4 wherein olefin production is optimized by employing fluid bedprimary stage conditions in the temperature range of about 425° C. to550° C., a pressure range of about 100° to 800° kPa and weight hourlyspace velocity range of about 0.5° to 3.0° based on ZSM-5 equivalentcatalyst and methanol equivalent in the primary stage feedstock.
 6. Theprocess of claim 4 wherein primary stage hydrocarbon effluent containsabout 1 to 10 wt. % ethene and about 10 to 60 wt. % C₃ -C₄ olefins. 7.The process of claim 4 wherein ethene is recovered from the primarystage effluent vapor stream by fractionation.
 8. The process of claim 1wherein the alkylation stage effluent is fractionated to provide aliquid hydrocarbon stream comprising C₇ to C₉ isoparaffins and wherein aportion of said C₇ to C₉ isoparaffinic fraction is recycled as sorbentliquid.
 9. The process of claim 1 wherein the secondary alkylation stagecomprises a liquid phase reaction catalyzed by HF.
 10. In the processfor producing alkylate hydrocarbons by catalytic reaction of isoparaffinwith lower olefin, the improvement which comprises:fractionating anolefinic feedstream containing ethene and C₃ + olefinic components bycontacting the olefinic feedstream in a sorption zone with a liquidhydrocarbon sorbent to selectively sorb C₃ + components; reacting C₃ +olefins with excess isoparaffin in a catalytic alkylation reactor toproduce C₇ + alkylate hydrocarbons; fractionating the alkylation reactoreffluent to provide a liquid hydrocarbon fraction rich in C₇ +,alkylate; introducing said C₇ + liquid hydrocarbon fraction to thesorption zone as lean sorbent; and recovering C₇ + alkylate product fromthe process.
 11. The process of claim 10 wherein the alkylation reactoreffluent is fractionated to provide an isoparaffin recycle stream forfurther conversion in the alkylation reactor.
 12. The process of claim 1wherein the alkylation catalyst comprises hydrofluoric acid and thereactor is operated at pressure of about 1500 to 3000 kPa to provide aliquid reaction phase.
 13. The process of claim 12 wherein theisoparaffin consists essentially of isobutane and the sorbent containsat least 75 weight percent of C₇ to C₉ paraffinic hydrocarbons.
 14. Theprocess of claim 10 wherein the olefinic feedstream is produced bycatalytic conversion of methanol over ZSM-5 catalyst at elevatedpressure and wherein ethene is recovered from the sorption zone and saidrecovered ethene is contacted with the catalyst for further conversionwith methanol.
 15. The process of claim 14 wherein the alkylationcatalyst comprises hydrofluoric acid and the reactor is operated atpressure of about 1500 to 3000 kPa to provide a liquid reaction phase.16. The process of claim 15 wherein the isoparaffin consists essentiallyof isobutane and the sorbent contains at least 75 weight percent of C₇to C₉ paraffinic hydrocarbons.